CN107438475B

CN107438475B - Method for energy-efficient recovery of carbon dioxide from an absorbent and apparatus suitable for operating the method

Info

Publication number: CN107438475B
Application number: CN201580074098.8A
Authority: CN
Inventors: R·芬德; J·F·波尔森
Original assignee: Consolidated Engineering Co Inc
Current assignee: Consolidated Engineering Co Inc
Priority date: 2014-12-11
Filing date: 2015-12-10
Publication date: 2021-10-08
Anticipated expiration: 2035-12-10
Also published as: EP3031511B1; ES2869573T3; CN107438475A; DK3229940T3; EA201791056A1; EP3229940A1; EP3229940B1; US20170361266A1; WO2016091266A1; US10730007B2; EP3031511A1

Abstract

The present invention relates to a process for recovering acid gas from an absorbent rich in gaseous acid gas, wherein the energy for separating the absorbent and acid in a stripper is reduced by recycling a heat transfer fluid from the stripper off-gas in an energy efficient manner. The invention further relates to an apparatus suitable for carrying out said method.

Description

Method for energy-efficient recovery of carbon dioxide from an absorbent and apparatus suitable for operating the method

The present invention relates to a method for recovering carbon dioxide from an absorbent enriched in gaseous carbon dioxide, wherein the energy for separating the absorbent and the carbon dioxide in a stripping column is reduced, and further to an apparatus for carrying out said method.

Background

Carbon dioxide recovery plants are widely used for cleaning and/or recovering carbon dioxide released from, for example, hydrocarbon combustion, fermentation and gas treatment. The recovered carbon dioxide may optionally be liquefied and sold as a final product or used in the production of a given plant.

In a typical absorber-regenerator system, carbon dioxide recovery is performed by introducing a gas into the absorber where the gas is contacted with a lean solvent containing the absorbent flowing downward in the absorber. The carbon dioxide is at least partially absorbed by the lean solvent, and the depleted gas exits the absorber for further processing or discharge. The carbon dioxide containing solvent is then treated, for example by stripping, to release carbon dioxide, and the carbon dioxide can be recovered or further purified. Conventional techniques that can be used to separately recover the absorbent and carbon dioxide include stripping. Absorber-regeneration systems typically allow for continuous operation to recover carbon dioxide.

In designing the process and defining the parameters that produce the desired high purity carbon dioxide while at the same time at the highest product yield, further downstream purification steps often include open and closed loop systems in which a lean stream of absorbent that will still contain residual carbon dioxide is further treated and recycled to extract even more carbon dioxide from the absorbent. The use of such a loop system thereby facilitates the recovery and reuse of feed streams such as absorbent and/or water to reduce costs and waste.

However, further processing for regenerating the absorbent or extracting more carbon dioxide from the absorbent also requires additional energy, for example for cooling, heating and pressurizing. In general, the energy consumption required per unit yield increases with the purity of the absorbent. That is, the energy consumption per unit yield required to recover the final residual carbon dioxide from the absorbent lean stream is higher compared to the recovery of the first unit, e.g., the carbon dioxide rich absorbent stream.

Several plants for recovering carbon dioxide have been described that improve the overall energy efficiency. US2013/0055756 describes one such recovery plant in which the lean absorbent from the regenerator column is recycled to the top of the regenerator column using conventional reboiling, and the mixed stripper off gas is compressed and condensed to recycle the solvent to the regenerator column for further recovery. Comprising an intermediate condensing heat exchange step wherein the lean absorbent and the mixed gas are heat exchanged. However, energy efficiency is in the heat pump system 6 located in and between the absorption and regeneration columns.

WO2008/063082 also describes a process for the regeneration of absorbed carbon dioxide, in which thermal energy is recovered from the carbon dioxide gas. The absorbed carbon dioxide stream is subjected to a stripping procedure, thereby producing a heated gaseous carbon dioxide-rich stream and a liquid carbon dioxide-depleted absorbent stream. The heated gaseous carbon dioxide-rich stream is passed through a plurality of compression steps with intermediate addition of a heat transfer fluid, and heat is recovered from the compressed gas using a heat exchanger. The recovered energy can be used to heat the liquid carbon dioxide-lean absorbent stream in the regeneration reboiler as heat or as part of the heat of the regenerator reboiler.

Thus, energy is recovered from the compressed heated carbon dioxide, and this heat can be used to reboil a portion of the lean absorbent to reduce the overall energy consumption of the carbon dioxide recovery process.

In US2014/0190351 a method is described for reducing the energy consumption of a carbon dioxide recovery process, more specifically for reducing the energy used in a stripping process, by using the thermal energy generated by the system for the stripper reboiler. This is addressed by providing low pressure steam to the stripper reboiler to provide stripping steam without degrading the solvent (absorbent). Condensate from saturated stripper gas can also be directed to the stripper reboiler for vaporization and use as stripping steam, thus reducing the need to supply make-up water.

In large facilities, even a small reduction in energy consumption per unit of carbon dioxide yield is of great economic benefit. Therefore, there is a continuing need to devise methods and parameters that enable the recovery of carbon dioxide with lower energy consumption. In addition to energy consumption, an equally important economic aspect is the reduction of consumption of other resources per unit of carbon dioxide yield, such as the amount of absorbent and/or water required for the process.

Also, it is always desirable to minimize the inputs to the process, such as feed water, absorbents, etc.

It is therefore an object of the present invention to further reduce the overall energy consumption of the carbon dioxide recovery process, as well as to reduce the consumption of additional resources, such as water.

Disclosure of Invention

The above object is achieved by aspects of the present invention, wherein a first aspect relates to a method for recovering carbon dioxide from a liquid carbon dioxide rich absorbent (L1), comprising the steps of:

a. providing a liquid acid gas-rich absorbent stream (L1) having absorbed acid gas therein,

b. separating carbon dioxide from the acid gas-rich absorbent stream (L1) in a stripping column (A2) using a gaseous stripping medium (G2) to provide an acid gas-containing gas stream (G1) and a liquid acid gas-lean absorbent stream (L2),

c. transferring heat from the acid gas-containing gas stream (G1) to a stream of heat transfer fluid (L4) to provide a cooled acid gas-containing stream (G3) and a heated stream (L4'),

d. separating the heated stream (L4') into recovered gaseous stripping medium (G4) and liquid heat carrier (L5), and

e. the recovered stripping medium (G4) is provided directly or indirectly to a stripping column (a 2).

Using this method, stripping steam recovered from the stripper off-gas is supplied to cover at least a portion of the same stripping process, thereby reducing the externally supplied heat. This provides an overall energy reduction wherein the external energy is replaced by steam energy generated by heat exchange of the stripper off-gas and the heat transfer fluid. In a particular embodiment, the acid gas is carbon dioxide.

In a preferred embodiment of the invention, the liquid heat carrier (L5) is used as at least a part of the heat transfer fluid (L4), whereby further recovery of the liquid used as heat transfer fluid is facilitated by providing a water/steam circuit. In addition, the heat carrier is kept in the stripping loop, thus minimizing the need to feed additional heat carrier (and eventually stripping medium) to the system, in addition to reducing heat supply.

In one embodiment, the absorbent is aqueous, e.g., the absorbent is an aqueous solution of the absorbent. Sterically hindered amines such as AMP, alkanolamines such as alkanolamines having 1-3 alkanol groups, each alkanol group having 1, 2, or 3 carbon atoms, and water are preferred. Examples of suitable alkanolamines are Monoethanolamine (MEA), Diethanolamine (DEA) and Triethanolamine (TEA), MEA being the absorbent of choice because it is recognized as inexpensive and has proven effective.

Using the process of the present invention, it has been demonstrated that the amount of energy required for this open loop for acid gas recovery can be reduced, and water saved.

That is, acid gas, e.g., carbon dioxide, from the feed stream and remaining acid gas in the lean absorbent is partially recovered by using vapors derived from the stripper process. This saves energy because less lean absorbent needs to be reboiled to provide sufficient stripping medium. Also, water from the stripping medium originating from the aqueous absorbent is recycled and reused. Thus, both the water and the energy required for providing a heat transfer fluid at a suitable temperature far above ambient temperature may be reduced.

In the present invention, the water and the remaining acid gas in the lean absorbent stream (L2) are reboiled by heating and returned to the second stripper (a 2).

The energy for heating the liquid acid gas-depleted absorbent stream (L2) to provide the gas/liquid heated acid gas-depleted absorbent stream (L2') is provided by external heating.

Thus, in one embodiment, the method further comprises the steps of:

f. heating the liquid acid gas-depleted absorbent stream (L2) to provide a gas/liquid heated acid gas-depleted absorbent (L2');

g. the gas/liquid heated acid gas-depleted absorbent (L2 ') is separated in a second separator (a7) to provide vaporized stripping medium (G2') and a recovered liquid absorbent stream L3.

The recovered liquid absorbent is preferably recycled to the absorber in a conventional manner.

In a preferred embodiment, the temperature of the recovered stripping medium (G4) is higher than the temperature of the gas/liquid heated acid gas-depleted absorbent (L2 '), and in another embodiment, the temperature of the mixture of the heated acid gas-depleted absorbent stream (L2') and the recovered stripping medium (G4) is from 100 ℃ to 130 ℃, preferably from 105 ℃ to 120 ℃, more preferably from 110 ℃ to 115 ℃.

The process of the present invention thus facilitates the use of alternative sources of thermal energy to provide the stripping medium. The source of thermal energy for stripping is provided by the recovered stripping medium (G4) originating from the stripping procedure itself, as it is recovered from the thermal energy of the acid gas-containing gas stream (G1) originating from the stripping procedure.

Thus, the process of the present invention facilitates the recovery of acid gas from a liquid acid gas-containing absorbent in an energy efficient manner.

In a particular embodiment, the recovered stripping medium (G4), optionally compressed, i.e. as described immediately below, is fed to the stripping column (a2) at a point below the feed point of the gaseous stripping medium (G2). The feeding of the recovered stripping medium (G4) or the compressed recovered stripping medium (G4') at a lower position provides a better recovery of carbon dioxide from the absorbent in the stripping column (a2), since the carbon dioxide content of the gaseous stripping medium is greater than the recovered compressed recovered stripping media (G4 and G4'), respectively.

In a preferred embodiment, the recovered stripping medium (G4) is compressed to provide a compressed recovered stripping medium (G4') prior to being provided to the stripping column (a 2). This ensures a smooth circulation of the compressed recovered stripping medium (G4') into the separator or stripper without the need for additional equipment. Thus, in addition to utilizing the heating capacity, recirculation of the stripping medium is facilitated easily by the pressure difference. It is contemplated that the pressure of the recovered stripping medium (G4) may be lower than, equal to, or higher than the pressure of the operating pressure in the stripping column (a 2).

The temperature of the compressed recovered stripping medium (G4') may be higher than the temperature of the gas/liquid heated acid gas-depleted absorbent (L2'). Because the two streams are mixed in one particular embodiment, the heat from the recovered stripping medium (G4) will enhance the evaporation of the stripping medium from the combined stream. This in turn lowers the temperature to which the absorbent-depleted stream (L2) has to be heated beforehand. Furthermore, the recovered stripping medium (G4) preferably has a pressure higher than or equal to the operating pressure of the separation by stripping in step b.

Thus, in one embodiment, the stripping medium (G2) is an evaporated stripping medium (G2'); compressed stripping medium (G4'), or both.

In one embodiment thereof, the compressed stripping medium (G4') and the vaporized stripping medium (G2') are mixed prior to being fed to the stripping column (a 2).

Thus, throughout the specification and claims, streams referred to as recovered stripping medium (G4) and compressed recovered stripping medium (G4') may be used interchangeably, and note whether or not compression is applied to the recovered stripping medium (G4).

Those skilled in the art will appreciate that certain temperature and pressure conditions must be maintained during the separation step to provide a processable stream that is returned to the stripping procedure. Thus, at about 1.4 bar, the minimum temperature of the stripping medium (G2) returned from the separation step should be about 100 ℃, preferably about 105 ℃ to 120 ℃, more preferably about 110 ℃ to 115 ℃.

In one embodiment, the ratio of mass flow rates between the stripping medium (G2) and the compressed recovered stripping medium (G4') in the mixing step is from 4:1 to 1:1, preferably from about 3.5:1 to 2.5:1, e.g., 3:1, provided that the two streams when mixed together have a temperature above 100 ℃, preferably above 105 ℃, more preferably above 110 ℃, or from 105 ℃ to 115 ℃, 110 ℃ to 115 ℃. However, other ranges of mixing would also be beneficial, for example as low as 5, 10 or 15% recycle of the total stripping steam (G4'). It is also contemplated that the pressure and temperature may be higher, such as temperatures up to 140 ℃, and pressures corresponding to saturated steam.

Thus, the recycle of recovered stripping medium (G4) replaces the external energy supplied to a6 that would otherwise be required to heat the liquid acid gas-lean absorbent stream (L2) to provide sufficient stripping medium. In one embodiment, the recovered stripping medium (G4) replaces external energy and the thermal energy is greater than 200kW, preferably greater than 300kW, as shown in the examples.

According to the invention, heat transfer from the acid gas-containing stream (G1) to the heat transfer fluid (L4) is facilitated by a heat exchanger. In a particular embodiment, the heat transfer in step c. is obtained by directly contacting a heat transfer fluid (L4) with an acid gas-containing gas stream (G1) to obtain a heated stream (L4') and a cooled acid gas-containing gas stream (G3). The direct contact of the two streams increases the utilization of the heat contained in the acid-containing gas stream (G1) since no heat is absorbed into the plant surfaces, which facilitates the exchange. Therefore, a direct contact cooler is a preferred unit for heat transfer in step c. Thus, heat transfer is conducted such that the temperature of the heated stream (L4') is approximately equal to the temperature of the acid gas-containing gas stream (G1), and the temperature of the cooled acid gas-containing stream (G3) is approximately equal to the temperature of the heat transfer fluid (L4). Complete heat transfer depending on the temperature difference can be obtained by adjusting the height of the direct contact cooler and/or the flow ratio of the streams. In general, one skilled in the art will be able to determine the appropriate cooler height and diameter for a given flow and temperature and/or flow for a given height and temperature.

In a specific embodiment where the sour gas is carbon dioxide, the flow rate of the carbon dioxide containing gas stream (G1) is 2400kg/h, the temperature is 102 ℃, the flow rate of the heat transfer fluid (L4) is 22000kg/h, and the temperature of the heat transfer fluid is 70 ℃, when the packing is pall rings (random packing), the heat transfer height of the direct contact cooler is about 4m and the diameter is about 0.7 m. This will provide complete transfer of heat from the carbon dioxide-containing gas stream to the heat transfer fluid, providing a cooled carbon dioxide-containing stream (G3) having a temperature of about 70 ℃ and a heated stream (L4') having a temperature of about 102 ℃.

In prior art processes, heat is typically recovered using indirect heat exchange, i.e. where the fluids are kept separately to prevent mixing. When direct contact is used, the heat transfer results in the heated stream (L4') containing dissolved acid gas remaining therein.

However, by direct contact cooling, a more efficient recovery of the heat present in the acid gas-containing stream (G1) may be achieved. Furthermore, the heated stream (L4 ') can be easily mixed with the gas/liquid heated acid gas-depleted absorbent stream (L2'). Thus, the majority of the water present in the acid gas-containing gas stream (G1) leaving the stripper is returned to the stripper. Thus, the overall water consumption in the acid gas recovery process is reduced, since the water used as stripping medium in the form of steam is recycled in a very efficient manner. Thus, only a small hydration dose, i.e. less than 1% of the feed flow (on a weight basis), is required. Furthermore, the dissolution of acid gas residue in the heat transfer fluid (L4) formed by direct contact heat transfer does not infer the total loss of acid gas. The dissolved acid gas in the recovered stripping medium (G4) is easily recovered in the stripping procedure, since the compressed recovered stripping medium (G4') is directly or indirectly mixed with the stripping medium (G2). Also, when direct cooling is used, the pressure drop in the unit will be lower than when indirect cooling is used, so the pressure of G3 will be higher and the final product G5 will eventually require less pressurization to provide it in a form suitable for further processing.

In one embodiment involving the use of direct contact heat exchange, prior to heat transfer step c, the acid gas-containing stream (G1) is compressed into a compressed acid gas-containing stream (G1').

In another embodiment using direct heat exchange, the separation of step d is provided by:

depressurizing the heated stream (L4') to provide a depressurized stream (L4 "); and

separating the depressurized stream (L4 ") by flashing in a first flash column (a9) to provide recovered stripping medium (G4) and liquid heat carrier (L5).

In yet another embodiment, the separating of step d.ii. provides a liquid stream (L4 "') that is subjected to the following steps, prior to providing the heat transfer fluid (L5):

further depressurizing the liquid stream (L4 ') to provide a second depressurized stream (L4 '), which is at a pressure lower than the pressure of the liquid stream (L4 '),

separating the second reduced-pressure stream (L4 "") in a third separation unit (a10) to provide a second gas (a) and a liquid heat carrier (L5);

recompressing the second gas (a) to provide recompressed second gas (b); and

d.iv. feeding the recompressed second gas (b) to a first flash column (a9) where it leaves the flash separation unit as part of the recovered stripping medium (G4).

By the steps provided by the process of the present invention, a means is provided to provide as much stripping medium as possible with a minimum energy input. Furthermore, by circulating water in the system in direct contact with the stream to be treated, the supply of both make-up water and energy required to reboil the acid gas-depleted absorbent (L2) is minimized.

In a particular embodiment, all steps d.i. to d.vi. are repeated, preferably 2, 3 or 4 times. In such embodiments, the repetition may be sequential and/or parallel, with more repetitions comprising steps d.i. to d.v. further increasing the amount of vapor.

In another embodiment, the heat transfer in step c.is provided by indirect heat exchange, wherein the heat transfer fluid (L4) is depressurized to a pressure lower than the pressure of the liquid heat carrier (L5) prior to heat transfer.

Depressurization prior to heat exchange has two advantages. First, it allows for better heat transfer so that the cooled acid-containing gas stream (G3) is cooler and therefore the heated stream (L4') is hotter than the pressure reduction after heat exchange. Secondly, it allows the heat recovery unit (A3) to be an indirect heat exchanger and the flash separation unit (a9) to be integrated into one unit. Thus, space and installation costs are saved.

Thus, in one specific embodiment using indirect heat exchange, the separation of step d. is provided by:

separating the heated stream (L4 ') by flashing in a first flash column (a9) to provide a recovered stripping medium (G4) and a liquid stream (L4 "');

further depressurizing the liquid stream (L4 "') to provide a second depressurized stream (L4" ") having a pressure lower than the pressure of the liquid stream (L4"');

recompressing the second gas (a) to provide recompressed second gas (b); and

v. feeding the recompressed second gas (b) to a first flash column (a9) where it leaves the flash separation unit as part of the recovered stripping medium (G4).

According to this particular embodiment, a further embodiment may be characterized in that all steps d.i. to d.v are repeated, preferably 2, 3 or 4 times. In such embodiments, the repetition may be sequential and/or parallel, with more repetitions comprising steps d.i. to d.v. further increasing the amount of vapor.

The use of indirect heat exchange will provide a cleaner stripper medium as there will be little, if any, acid gas dissolved in it. Thus, in this embodiment, the full heat transfer is to exchange the stripping medium for more acid gas-lean, so eventually more acid gas-lean absorbent is recycled to the absorption column.

The following features and variables are all referred to conventional procedures and can be applied equally to all embodiments, i.e. whether direct or indirect heat transfer. Thus, in one embodiment, any one or more of L4 ', L4 ", L4"', L4 "", preferably L4 "and/or L4" ", is heated by a heat source, preferably a low value heat source. Supplying additional heat to any of these streams will provide increased vapor pressure for the particular stream, thus increasing the mass flow of G4. The heat may be taken from any suitable location within the process itself or from an external heat source. One skilled in the art will appreciate that when in excess, heat is present in any given system.

One embodiment of the present invention comprises the steps of: optionally cooling the cooled acid gas-containing stream (G3), and separating the optionally cooled stream into an acid gas product stream (G5) and a second liquid stream (L6), and optionally recycling the second liquid stream (L6) to the stripper column (a 2).

In yet another variant of the inventive process, the liquid acid gas-depleted absorbent stream (L2) is heated in a third heat exchanger (a6) and separated to provide a gaseous stripping medium (G2) and a recovered liquid absorbent stream (L3), wherein the heat transfer fluid (L4) of step c is the recovered liquid absorbent stream (L3), and wherein the heat recovery unit (A3) for heat transfer is an indirect heat exchanger (A3).

In this way, the heat transfer fluid is lean absorbent. With this embodiment, heat will be transferred at a reduced installation cost.

In a particular embodiment of this variant, the recovered liquid absorbent stream (L3) is depressurized in a fourth depressurization unit (a18) to provide a depressurized recovered liquid absorbent stream (L3'), which may be a depressurization valve or a flow control valve. This stream is heated in a heat recovery unit (A3) to provide a heated recovered liquid absorbent stream (L3 "), it is further contemplated that the heated recovered liquid absorbent stream (L3") is separated in a fourth separator (a19) to provide absorbent and recovered stripping medium (G4), which may be pressurized in a second pressurization unit (a12) to provide compressed recovered stripping medium (G4'), which is provided to the stripping column (a 2).

In another aspect, the invention relates to an apparatus for removing sour gas from a feed gas, the apparatus comprising a stripper column (a2), the stripper column (a2) having a gas inlet through which a stripping medium is fed and a liquid inlet through which a liquid sour gas-rich absorbent is fed, the stripper column (a2) having a gas outlet and a liquid outlet, the gas outlet being connected to a heat recovery unit (A3), the heat recovery unit (A3) additionally having a heat transfer fluid inlet, a product gas outlet and a second liquid outlet, the second liquid outlet being directly or indirectly connected to the stripper column (a 2). In another embodiment, the second liquid outlet of the heat recovery unit (A3) is connected to a first depressurization unit (A8), which first depressurization unit (A8) is then directly or indirectly connected to the stripper column (A3).

In one embodiment, the apparatus comprises a second separator (a7), wherein separator (a7) has a liquid inlet connected to a liquid outlet of stripper column (a2) through a third heat exchanger (a6), and said separator (a7) additionally has a liquid outlet and a gas outlet, said gas outlet being connected directly or indirectly to a gas inlet of stripper column (a 2).

In another embodiment, the first depressurization unit (A8) is connected to a flash separation unit (a9), the flash separation unit (a9) further having a gas inlet, a liquid outlet and a gas outlet, the gas outlet being directly or indirectly connected to the stripper column (a2), and the liquid outlet being directly or indirectly connected to the heat recovery unit (A3).

In one embodiment, the flash separation unit (a9) is connected to a second pressure boosting unit (a12), which second pressure boosting unit (a12) is connected directly or indirectly to the stripper column (a 2).

In another embodiment, the liquid outlet of the flash separation unit (a9) is indirectly connected to the heat recovery unit (A3) by way of the liquid outlet being connected to a second pressure reduction unit (a13), said second pressure reduction unit (a13) being directly or indirectly connected to the liquid inlet of the heat transfer fluid inlet of the heat recovery unit (A3).

In one embodiment, the second pressure reduction unit (a13) is connected to a third separator (a10), said third separator (a10) additionally having a gas outlet and a liquid outlet, said gas outlet being connected to a first pressure increasing unit (a11), said first pressure increasing unit (a11) being connected to the gas inlet of the flash separation unit (a9), wherein the liquid outlet of the second separator is connected to the liquid inlet of the heat transfer fluid inlet of the heat recovery unit (A3), optionally through a pump and/or mixer unit (a 14).

Thus, in a preferred embodiment of the present invention, the first depressurization unit (A8) is indirectly connected to the stripper column (a2), wherein the first depressurization unit (A8) is connected to the flash separation unit (a9), said flash separation unit (a9) further having a gas inlet, a liquid outlet and a gas outlet, said gas outlet being connected to the second depressurization unit (a12), said second depressurization unit (a12) being connected to the stripper column (a2), the liquid outlet being connected to the second depressurization unit (a13), said second depressurization unit (a13) being connected to the third separator (a10), said third separator (a10) further having a gas outlet and a liquid outlet, said gas outlet being connected to the first depressurization unit (a11), said first depressurization unit (a11) being connected to the gas inlet of the flash separation unit (a9), wherein the liquid outlet of the third separator (a10) is optionally connected to a14 via a pump and/or a mixer unit (a14), is connected to the liquid inlet of the heat transfer fluid inlet of the heat recovery unit (a 3).

In another embodiment, the heat recovery unit (a3) directly contacts the heat exchanger.

It is contemplated that one or both of the pressure reducing units (A8) and (a12) are pressure reducing valves and/or one or both of the pressure increasing units (a11 and a12) are vacuum pumps (compressors).

In one embodiment, the liquid outlet of the heat recovery unit (A3) is indirectly connected to the second separator (a7) in such a way that the connection of the liquid outlet of the heat recovery unit (A3) is in fluid communication with the stripper column (a2) via the second separator (a 7).

The apparatus is designed to run the method of the invention.

Drawings

Examples of embodiments of the present invention are described in more detail below with reference to the accompanying drawings, in which

Figure 1 shows a process and apparatus for acid gas recovery according to one of the simplest aspects of the invention when direct contact heat exchange is used.

Figure 2 shows a method and apparatus using indirect contact heat exchange.

Figure 3 shows a process and apparatus using direct contact heat exchange and a subsequent flash separation and separation loop.

Fig. 4 shows the process and apparatus of fig. 3 wherein the acid gas-containing gas stream is additionally pressurized prior to heat exchange.

Figure 5 shows a process and apparatus using indirect heat exchange followed by a flash separation and separation loop where the direction of the recovered stripping medium is not specified.

Figure 6 shows a process and apparatus using indirect heat exchange wherein the compressed recovered stripping medium is returned directly to the stripping column.

Figure 7 shows a process and apparatus using indirect heat exchange wherein compressed recovered stripping medium is returned to the second separator.

Figure 8 shows a process and apparatus in which a plurality of additional direct contact heat exchange heat transfer circuits and separations are used.

Figure 9 shows a process and apparatus wherein the heat transfer fluid is a recycled liquid absorbent.

Fluid and influent/effluent

Throughout the specification and claims, the streams and influent/effluent are represented by: a mild liquid acid gas/carbon dioxide rich absorbent L0; liquid acid gas/carbon dioxide rich absorbent L1; a liquid acid gas/carbon dioxide-depleted absorbent stream L2; gas/liquid heated acid gas/carbon dioxide lean absorbent L2'; recovered liquid absorbent stream L3; a depressurized recovered liquid absorbent stream L3'; a heated recovered liquid absorbent stream L3 "; heat transfer fluid L4; heated stream L4'; reduced pressure stream L4 "; liquid stream L4' ″; a second depressurized stream L4 ""; liquid heat carrier L5; a second liquid stream L6; acid gas/carbon dioxide containing gas stream G1; a compressed acid gas/carbon dioxide-containing gas stream G1'; stripping medium G2; evaporated stripping medium G2'; a cooled acid gas/carbon dioxide-containing stream G3; recovered stripping medium G4; compressed recovered stripping medium G4'; acid gas/carbon dioxide product stream G5, second gas a; a compressed second gas b.

Component part

Throughout the specification and claims, the components of the apparatus are represented by: a first heat exchanger a 1; stripper a 2; heat recovery unit a 3; a second heat exchanger a 4; a first separator a 5; the third heat exchanger a 6; a second separator a 7; a first decompression unit A8; flash separation unit a 9; the third separator a 10; the first pressure increasing unit a 11; the second pressure increasing unit a 12; a second decompression unit a 13; mixer unit a 14; the third pressurizing unit a 15; cell A16; a third decompression unit a 17; a fourth pressure reduction unit a18, a fourth separator a 19.

Detailed Description

The illustrations appended to this description should be understood to be a section of a larger facility. All features and variables described herein for each embodiment apply equally to all embodiments. Thus, detailed features relating to the method may equally be applied to the apparatus and vice versa. For the sake of brevity, no auxiliary devices are included in the figures. However, the type and location of the device will be readily understood by those skilled in the art. As examples of auxiliary devices, mention may be made of liquid pumps, valves, condensers for condensing water and/or chemical absorbents from the exhaust gases, devices for replenishing water and/or absorbents, etc. In the following detailed description, the invention will be described with reference to carbon dioxide as an example of sour gas. The invention should not be limited thereto and all embodiments apply equally to acid gases in general, for example CO₂、H₂S、SO₂And the like.

Throughout the specification and claims, the terms "rich", "lean" and "depleted" refer to the amount of, for example, carbon dioxide or absorbent contained in a particular stream, which name may be used to distinguish different streams formed by different separation steps, and should be interpreted relative to each other in a particular separation step.

Referring now to fig. 1, a process according to the first aspect of the invention is shown below in one of the most general forms, the absorbent being described as an aqueous solution, but the invention should not be limited thereto.

The schematic should be interpreted as being the downstream region in the absorber column after absorption of carbon dioxide from the gas source. The absorption of carbon dioxide is well known in the art. A method of recovering carbon dioxide from a carbon dioxide rich absorbent comprising the steps of: a liquid carbon dioxide-rich absorbent stream L1 is provided having absorbed gaseous carbon dioxide.

The carbon dioxide absorbed in the liquid carbon dioxide rich absorbent L1 may originate from any kind of carbon dioxide source. The carbon dioxide source may originate, for example, from the combustion of fossil fuels, from flue gases, from the production of synthesis gas, or from a production line filled with carbonated beverages. Preferably flue gas.

The absorbent preferably absorbs chemically. Thus, the liquid carbon dioxide rich absorbent L1 comprises a suitable absorbent for carbon dioxide or other acid gases. The absorbent for absorbing gaseous carbon dioxide may be any solvent known to be capable of absorbing carbon dioxide and/or acid gases. By way of example, mention may be made of solutions of alkanolamines, more preferably alkanolamines having from 1 to 3 alkanol residues, more preferably alkanol residues having 1, 2 or 3 carbon atoms, in aqueous solution. Examples of suitable alkanolamines are Monoethanolamine (MEA), Diethanolamine (DEA) and Triethanolamine (TEA), MEA being the absorbent of choice because it is recognized as inexpensive and effective. The concentration of the absorbent is typically 5-40% aqueous solution, one example is monoethanolamine, which is a 35% aqueous solution of MEA.

After absorption, the rich absorbent may be heated and/or pressurized before being fed to stripper a 2. It is known that carbon dioxide is separated from alkanolamine absorbent by heating, preferably to a temperature of 90 ℃ and above, preferably above 90-110 ℃, more preferably 95 ℃ or above, 100 ℃ or above, e.g. 104 ℃ to 106 ℃. The pressure of the liquid carbon dioxide rich absorbent L1 may be increased above atmospheric pressure, for example 1-3 bar, or to a pressure above the pressure of the stream leaving the absorption column, which is often operated at ambient pressure. The stripping process itself may typically be operated at slightly above atmospheric pressure (above the operating pressure of the absorber column), for example from 1.2 to 2.6 bar, more preferably from 1.2 to 1.6 bar, for example 1.4 bar. Other pressures are contemplated within the context of the present invention.

The liquid carbon dioxide rich absorbent L1 was stripped using a stripping medium G2, which mainly contained steam, i.e. water vapor, when the absorbent contained water. Advantageously, according to the invention, the stripping medium is a partially recycled water coming from the stripping procedure itself.

The stripping medium is stream G2, which is substantially absorbent-free and contains water vapor, with a low carbon dioxide content. In the context of the present invention, very low is less than 10 mol%, more preferably less than 5 mol%. In one embodiment, the temperature of the stripping medium G2 is higher than the temperature of the heated liquid carbon dioxide rich absorbent L1, more specifically higher than 100 ℃, preferably from 105 ℃ to 120 ℃, preferably about 115 ℃. This will more efficiently strip carbon dioxide from the absorbent.

In stripper a2, carbon dioxide is stripped from the absorbent to provide a carbon dioxide-containing gas stream G1 and a liquid carbon dioxide-depleted absorbent stream L2.

The carbon dioxide containing gas stream G1 is provided at temperature and pressure conditions above the vaporization conditions of water. Thus, the carbon dioxide containing gas stream G1 is a mixture of carbon dioxide and water vapor.

Carbon dioxide-depleted absorbent L2 typically exits the lower portion of stripping unit a 2. However, it is contemplated that the stream may be withdrawn from any suitable location of the column.

In the context of the present invention, the term "lean" is intended to mean a stream containing an amount of absorbed carbon dioxide which is lower than the amount of carbon dioxide in the liquid carbon dioxide rich absorbent L1. Thus, the liquid carbon dioxide-depleted absorbent stream L2 comprises less than 10 mol% carbon dioxide, typically less than 5 mol% carbon dioxide.

The liquid carbon dioxide-depleted absorbent stream L2 is then heated by indirect heat exchange in a third heat exchanger a6 to provide a gas/liquid heated carbon dioxide-depleted absorbent L2', which is a gas/liquid mixture. Reboiling is typically achieved using a closed steam-generating loop as is common in the art, but other means are also contemplated. In the embodiment shown, the gas/liquid mixture is further separated in a second separator a7, which provides vaporized stripping medium G2' and recovered liquid absorbent L3. The vaporized stripping medium G2 was fed as stripping medium G2' to stripper a2 and the recovered liquid absorbent L3 was returned to the absorber, most often in heat exchange relationship with liquid carbon dioxide rich absorbent L1 as is common in the art.

A portion of the thermal energy contained in the carbon dioxide containing gas stream G1 is transferred to the heat transfer fluid L4, thereby providing a cooled carbon dioxide containing gas stream G3 and a heated stream L4'.

Heat transfer is carried out in heat recovery unit a 3. Preferably, the heat exchange is carried out by direct contact, wherein the heat transfer fluid L4 and the carbon dioxide containing gas stream G1 are in physical contact with each other.

By using direct contact, heat exchange will be more efficient, and can reach almost 100%, depending on the size and/or flow rate of the individual streams. In addition, both the water condensed from the carbon dioxide containing gas stream G1 and the water from the heat transfer fluid will be mixed and can be used for recycle to the stripper and heat transfer, respectively.

In that way, it is ensured that an absolutely minimum water/fluid supply is required in the process, which will save the overall costs of the process.

Consequently, the temperature of the heat transfer fluid L4 is necessarily lower than the temperature of the carbon dioxide-containing gas stream G1. In one embodiment, the temperature of the carbon dioxide containing gas stream G1 is between 90 ℃ and 115 ℃ and the heat transfer fluid L4 is between 65 ℃ and 80 ℃.

Thus, in an exemplary embodiment, the heat exchanger is a direct contact cooler, and the temperature of heat transfer fluid L4 is about 70 ℃, the temperature of carbon dioxide-containing gas stream G1 is typically 102 ℃, and the mass flow ratio of the streams G1: L4 is about 1:9(kg/h)/(kg/h), the heat transfer height of the direct contact cooler is about 4m and the diameter is about 0.7m when the packing is pall rings (random packing).

With these correlation values, full heat transfer will occur such that the temperature of the heated heat transfer fluid is about the same as the temperature of carbon dioxide-containing gas stream G1, and the temperature of the cooled carbon dioxide-containing stream G3 is the same as the heat transfer fluid L4. It is contemplated that the temperature, flow rate and size may vary. Depending on the choice of the specific parameters, the person skilled in the art will be able to determine the remaining parameters, for example using any simulation program suitable for thermodynamic calculations; such procedures are well known in the art.

This step thus provides for the recovery of thermal energy from the carbon dioxide containing gas stream G1 to the heat transfer fluid L4. The cooling also results in condensation of water vapour present in the carbon dioxide containing gas G1, which provides a heated stream L4' comprising condensed water vapour from the carbon dioxide containing gas stream G1, and the heat transfer fluid L4 (now heated).

This cooling of the carbon dioxide containing gas stream G1 provides water from the process which is further treated and recycled to the stripping step as a stripping medium and in particular embodiments also as a heat transfer fluid.

Thus, the major portion of the heat transfer fluid L4 preferably originates from the process itself in the stripping medium loop, which is condensed and recycled, and is used as at least a portion of the heat transfer fluid.

In another specific embodiment, a portion of heat transfer fluid L4 originates from the absorption step prior to the stripping step. If the absorption step produces excess heat, the absorber may have a cooling device such as a reflux condenser. Its function is to reduce the loss of absorbent and evaporation of water from the absorber. The cooled condensate water and absorbent and the condensate may be used as, or as part of, heat transfer fluid L4. Thus, as the absorption generates heat, more water is collected internally, thereby allowing more steam to be generated for the stripping step.

Thus, the heat transfer fluid L4 or a portion of the heat transfer fluid may be externally supplied, originating from an absorption process prior to the process of the present invention, or a combination of all of the above.

The heated stream L4' is subjected to a flash separation step in a flash separation unit a9, which provides recovered stripping medium G4 and condensed stripping medium L5.

The cooled carbon dioxide-containing stream G3 described above is further heat exchanged and separated in a first separator a5 to provide a carbon dioxide product stream G5 and a second liquid stream L6, this further cooling of the cooled carbon dioxide-containing stream G3 ensuring that even more water (fluid) is removed from the carbon dioxide gas. At this point, the gas will contain less than 5 mol% water. If desired, the carbon dioxide product stream G5 may be subjected to final purification steps such as condensation, distillation, adsorption, or combination.

In the following description of the figures, all the flows and steps are the same as they are shown in FIG. 1.

Referring now to fig. 2, another embodiment is shown wherein heat exchange between the carbon dioxide containing gas stream G1 and the heat transfer fluid L4 is by indirect cooling. Thus, the main difference between the embodiments shown in fig. 1 and 2 is the type of heat transfer used in step c. In fig. 2 indirect heat exchange is used. Thus, no mixing of the two streams G1 and L4 takes place.

In this embodiment, a third pressure reduction unit a17 is shown prior to heat transfer by the heat recovery unit A3 to reduce the pressure of the heat transfer fluid L4. The liquid heat carrier L5 recovered from flash separation unit a9 was recycled to mixing unit a14 to mix with makeup water and the combined stream was depressurized to provide heat transfer fluid L4. It is also contemplated that the heat recovery unit (A3) and the flash separation unit (a9) are integrated into one unit.

It is also contemplated (as shown in fig. 2) that the second liquid L6 is split in a first unit a16_1, with a portion returned to stripper column a2 for further purification, and a second portion mixed with the liquid heat carrier in a second unit a16_2 before entering mixer unit a 14.

Prior to being fed to stripper a2, the mild liquid carbon dioxide-rich absorbent L0 may be heated in heat exchanger a1 to provide a liquid carbon dioxide-rich absorbent L1 at a temperature higher than the temperature of the mild liquid carbon dioxide-rich absorbent stream L0. Thus, the temperature of the liquid carbon dioxide rich absorbent stream L1 is preferably in the range of 90 ℃ to 110 ℃, more preferably in the range of 103 ℃ to 105 ℃, most preferably 104 ℃, as is known for separating carbon dioxide from alkanolamine absorbents.

In general, the heat exchange in exchanger a1 is typically a heat exchange between the hotter lean absorbent from stripper L3 and the milder rich absorbent L0 from the absorber. It is also contemplated (not shown) that this stream is drawn off by suction, for example by a vacuum pump, alternatively or additionally to heat exchange with the hotter lean absorbent from stripper L3. Suction improves operating efficiency over heating. A variation of this embodiment is further shown in fig. 9 and described below.

The embodiment shown in fig. 3 also has a direct contact cooler. In this embodiment, a heat transfer loop is introduced. The heated stream L4' provided by the heat transfer step c is subsequently depressurized into a depressurized stream L4 ". The pressure reduction is preferably obtained by means of a valve, more specifically a pressure reducing valve A8. The depressurized stream L4 "is a gas/liquid mixture. This stream is depressurized to a pressure lower than the pressure of the heated stream L4', preferably to a pressure lower than atmospheric pressure to provide a liquid/gas mixture, i.e. lower than 1 bar/1 atm. Typically, the pressure is reduced to about half the pressure of the prior stream, e.g., from about 1.4 bar to about 0.7 bar, etc.

A heat transfer circuit is included to provide liquid heat carrier L5. Flash separation of the depressurized stream L4 "provides a liquid stream L4'" and recovered stripping medium G4. The liquid stream L4 "' of condensed water is further depressurized in a second depressurization unit a13 to provide a second depressurized stream L4" ", which stream is a gas/liquid mixture. The pressure is likewise preferably halved, for example to 0.4 (when the first pressure drops to 0.7 bar). The second reduced pressure stream L4 "" is then separated in third separator a10 to provide liquid heat carrier L5, which consists essentially of water, which is recycled and mixed with makeup water to provide heat transfer fluid L4.

The pressure of heat transfer fluid L4 may be increased to a pressure corresponding to the stripping system, which is typically about 1-1.5 bar, preferably 1.4 bar, prior to mixing. Alternatively, when cooling is by indirect cooling, the pressure may be adjusted to a lower pressure.

The separation in the third separator a10 also provides a second gas a of water vapour. This stream is recompressed to the operating pressure of flash separation unit a9 to provide compressed second gas b, which is supplied to flash separation unit a9, where it is flashed with reduced pressure stream L4 ", which provides recovered stripping medium G4.

Thus, with this cycle, the amount of steam extracted from the liquid phase intended to be used as stripping medium is increased in a simple manner and with a minimum supply of energy.

The recovered stripping medium G4 provided by the flash separation is then, in the embodiment shown, recompressed in the second pressurizing unit a12, producing compressed recovered stripping medium G4'. The second pressurizing unit may be vacuum grass seed, but suitable alternatives are contemplated. The compressed recovered stripping medium G4' has a temperature significantly higher than the temperature of the recovered stripping medium G4. The heat originates from recompression. The compressed recovered stripping medium G4' is then mixed with a gas/liquid heated carbon dioxide-lean absorbent L2 ' in a second separator a7 to provide an evaporated stripping medium G2 ' and fed to a stripper a 2. It is also contemplated that the compressed recovered stripping medium G4 'is fed directly to stripper a2 or mixed with vaporized stripping medium G2' prior to entering stripper a 2.

The embodiment shown in fig. 4 is the same as the embodiment of fig. 3, except that the carbon dioxide containing gas stream G1 is compressed using a third pressure increasing unit a15 before entering the heat recovery unit A3 to provide a compressed carbon dioxide containing gas stream G1'.

Providing a pressurizing step prior to heat exchange may be used in combination with both direct contact heat exchangers and indirect heat exchangers (not shown). When a compression step is included at this point, the heat transfer in step c. will be more efficient and the gas fraction of the depressurized stream L4 "will provide a larger vapor fraction before separation in separator a 9.

According to the present invention, several embodiments are provided to use recovered stripping medium G4. In particular, the generality is shown in fig. 5-7, where the recovered stripping medium G4 can be provided directly or indirectly to a stripper a 2. Thus, in the embodiment shown in fig. 5, recovered stripping medium G4 is not explicitly connected to second separator a7, or stripper a2, and those skilled in the art will recognize that to provide this most versatile embodiment, the energy contained in recovered stripping medium G4 can be redirected to the most efficient point in accordance with the present invention. However, in the embodiment shown in fig. 6, recovered stripping medium G4 is compressed to provide a compressed recovered stripping medium stream G4', which is returned directly to stripping column a 2. Although not shown in fig. 6, it will also be appreciated that the compressed recovered stripping medium G4' is fed to stripper column a2 at a location below the location where the gaseous stripping medium G2 is fed. It will also be appreciated that in the case of a pressure sufficiently high by pressurization or gravity or a flow facilitated by a liquid pump, the recovered stripping medium is fed directly to the stripping column without prior compression. In the embodiment shown in fig. 7, compressed recovered stripping medium G4' is connected to second separator a 7. The compressed recovered stripping medium G4 'is separated in a second separator a7 together with a gas/liquid heated carbon dioxide-depleted absorbent stream L2' to provide a stripper medium G2 as a gas phase and a recovered liquid absorbent stream L3 which is returned upstream of the absorption procedure.

By recompressing the recovered stripping medium G4, the stream is heated and a large flow of the stripper medium can be provided. Also, in a preferred embodiment, the pressure of the compressed recovered stripping medium G4' corresponds to the operating pressure of the stripping column, whereby the recycle stream can be used directly without further treatment.

In the embodiment shown, heat transfer fluid L4 is derived from the recycle of water from the stripper and water from the flue gas/evaporated water from the absorbent. Thus, the heat transfer fluid L4 is essentially a recycle stream from the process itself. In such an embodiment, there will be a very low amount of externally supplied make-up water, and this portion will constitute less than 5% (mol/mol), more preferably less than 3% (mol/mol), even more preferably less than 1% (mol/mol) of the water mass flow in the system.

In another embodiment of the process shown in fig. 8, the heat transfer circuit introduced in fig. 3 was further developed to include multiple heat transfer circuits, in this particular embodiment two circuits connected in parallel are shown. In the embodiment shown, the heat transfer is direct heat exchange, but indirect heat exchange may equally be used.

Therefore, [ n ] in the figure indicates a loop number other than the first loop. When there is only one loop, the numbering is omitted. Thus, in the context of the present invention, n will be 2 or an integer greater than 2, such as 2, 3, 4, 5, 6, etc.

Thus, in this embodiment, cooled carbon dioxide-containing stream G3 exits direct contact cooler A3 and enters second direct contact cooler A3[2], wherein this stream is further cooled with a recycle stream of second liquid heat carrier L5[2], which provides second heated stream L4' [2] and second cooled carbon dioxide-containing stream G3[2 ]. Similar to the first loop, the second heated stream L4' [2] is depressurized (A8[2]), and separated (A9[2]) into a second recovered stripping medium G4[2] and a second liquid heat carrier L5[2 ]. The second recovered stripping medium G4[2] is compressed and mixed with recovered stripping medium G4. The combined stream is then fed to a stripper column, optionally before compression and separation, as shown in fig. 8.

In the same manner as described above, an additional heat transfer circuit may be added, in which case the next input gas to the third contact cooler (A8[3]) will be the second cooled carbon dioxide-containing stream G3[2], and so on.

The heated stream L4' undergoes substantially the same steps as described in fig. 1, i.e., it is depressurized to provide a depressurized stream L4 "which is separated in flash separation unit a9 to provide recovered stripping medium G4 and liquid heat carrier L5.

In the embodiment shown, the liquid heat carrier is recycled and mixed with a makeup stream to provide heat transfer fluid L4, and recovered stripping medium G4 is mixed with compressed second recovered stripping medium G4 '[ 2], with the combined portions being compressed in compressor a12 to provide compressed recovered stripping medium G4', which as shown is fed to second separator a 7. However, with this embodiment, it is also contemplated that the compressed recovered stripping medium G4' is fed directly to the stripping column or mixed with the stripping medium G2 prior to being fed to the stripping column.

Like the parallel circuits, it is also advantageous to have series circuits. Referring to fig. 3, in the case of a more serial circuit, the liquid stream (liquid heat carrier L5) leaving the third separator a10 will be depressurized to provide a gas/liquid mixture, which stream will be separated in another separator (a10[2]), where the liquid will correspond to heat transfer liquid L5, which is recycled as part of the heat transfer fluid, and the gas will be compressed in a compressor (a11[2]), fed to the third separator a10, and from there follow the route described in fig. 3.

In the same manner as described above, it is contemplated that multiple circuits may be inserted in parallel or in series. Preferably there are 1, 2 or 3 circuits in parallel and/or in series, for example two parallel circuits and 1, 2 or 3 series circuits.

By including n parallel loops, more fluid (water) is condensed from product gas G3[ n ] (where n is an integer such as 2, 3, 4, 5, 6, etc.), providing a cleaner product stream, avoiding some downstream purification steps, and providing more stripping steam for stripping. The fluid/water in the product stream is conventionally provided to the stripper as a liquid and is therefore not used as a stripping medium. By including more circuits in series, more steam is provided to the stripping.

Referring to fig. 9, another embodiment of the invention is shown wherein the acid gas-containing gas stream G1 leaving stripper a2 is fed to a heat recovery unit A3, which in the embodiment shown is an indirect heat exchanger. The recovered liquid absorbent stream L3 was depressurized in a fourth depressurization unit a18 (shown as a valve) to provide a depressurized recovered liquid absorbent L3' which was then passed through a heat recovery unit A3. Heat from the acid gas-containing gas stream G1 is transferred to the depressurized recovered liquid absorbent stream L3' to provide a heated recovered liquid absorbent stream L3 ", which is a gas/liquid mixture. The heated recovered liquid absorbent stream L3 "is separated in a fourth separator a19 to provide recovered absorbent and recovered stripping medium G4. The recovered stripping medium is compressed (e.g. in the second pressure increasing unit a12) as compressed recovered stripping medium G4' before being fed to the stripper column a 2.

The preferred values may vary depending on the ratio of the cost of the additional equipment to the cost of providing additional stripping steam and the reduction in the third heat exchanger a 6.

It has been revealed that when the solvent is aqueous and when the temperature of the recovered stripping medium (G4) is 70 ℃, the partial pressure of water is 0.31 bar (absolute), further temperature reduction will not produce any significant increase in heat recovery due to condensation of water, and therefore no significant effect is seen by including more loops.

Another aspect of the invention relates to an apparatus for recovering carbon dioxide from a liquid carbon dioxide-rich absorbent, the apparatus comprising a stripper a2, the stripper a2 having a gas inlet through which a stripping medium (G2) can be fed and a liquid inlet through which a liquid carbon dioxide-rich absorbent (L1) can be fed, the gas outlet being connected to a heat recovery unit A3, the heat recovery unit A3 additionally having a heat transfer fluid inlet, a product gas outlet and a second liquid outlet, the second liquid outlet being directly or indirectly connected to the stripper a 2. In another embodiment, the second liquid outlet of the heat recovery unit A3 is connected to a first depressurization unit A8, which is then connected directly or indirectly to a stripping column (A3). In one embodiment, the apparatus comprises a second separator a7, wherein the second separator a7 has an additional liquid inlet connected to the liquid outlet of stripper a2 through a third heat exchanger a6, a liquid outlet, and a gas outlet directly connected to the gas inlet of stripper a 2.

In another embodiment of the plant, the depressurization unit A8 is indirectly connected to the stripper A2, the first depressurization unit A8 is connected to the flash separation unit A9, the flash separation unit a9 further has a gas inlet, a liquid outlet and a gas outlet, the gas outlet being connected to a second pressure increasing unit a12, the second pressurizing unit A12 is connected to the second separator A7, the liquid outlet is connected to the second depressurizing unit A13, the second pressure reduction unit A13 is connected to a third separator A10, the third separator A10 additionally having a gas outlet and a liquid outlet, the gas outlet is connected to a first pressure increasing unit A11, the first pressure increasing unit A11 is connected to the gas inlet of a flash separation unit A9, wherein the liquid outlet of the second separator a7 is connected to the liquid inlet of the heat transfer fluid inlet of the heat recovery unit A3, optionally by a pump and/or an optional mixer unit a 14.

In another embodiment, the heat recovery unit a3 is a direct contact heat exchanger.

It is contemplated that one or both of the pressure reducing units A8 and a12 are pressure reducing valves and/or one or both of the pressure increasing units a11 and a12 are operated with vacuum pumps.

The heat exchanger used in the present invention may in principle be of any type. It is within the ability of one skilled in the art to select an appropriate heat exchanger based on the estimated flow rates of the absorbent, gas and receiving medium and the estimated amount of thermal energy to be transferred to provide a better temperature in the absorber and to recover heat to improve the energy economy of the overall plant.

The plant further comprises a stripper a 2. The stripping column to be used in the plant may be any packed column known in the art. Examples of suitable stripping columns are columns comprising internals or mass transfer elements such as trays or random or structured packing. Typically, the heated carbon dioxide-containing gas stream exits the top of the stripper and the liquid carbon dioxide-depleted absorbent stream exits the bottom of the stripper.

The purified gaseous carbon dioxide stream leaves the top of the heat recovery unit A3 after cooling and is then separated in a first separator a5 to remove the final small amount of water in the stream. Purified carbon dioxide leaves the separator as a gaseous product and the condensed water stream returning to the stripping column leaves the bottom of condensing unit a 5.

The stripper a2 may further have an additional inlet for make-up water fed at the top of the stripper and an additional inlet for condensed carbon dioxide depleted water recovered in the first separator a 5.

The plant further comprises a heat recovery unit a3 for cooling the carbon dioxide containing gas stream. The heat recovery unit may be a direct or indirect contact cooler, preferably a direct contact cooler. Based on the mass flow rate and the temperature and pressure conditions, the skilled person will be able to determine the height of the direct contact heat exchanger required for cooling the carbon dioxide containing gas stream G1.

In one embodiment, the apparatus further comprises one or more flash units and one or more separators, as is well known in the art. The flash unit of the present invention may be a simple vapor-liquid separation tank or a vapor/liquid separator. In particular, the separators a5, a7 and a10 are simple liquid/gas separators, also known as knock-out drums or flash tanks, having only one inlet and thus gas and liquid outlets. The flash unit a9 is preferably a flash separation unit, but may also be a simple knock out drum if the stream is clean enough.

The apparatus further comprises one or more pressure increasing units for compressing the heated cooling fluid into a vapor. The pressurizing unit may be a compressor, a steam ejector, a blower, or the like. In the present invention, since the preferred stream is steam, the preferred booster unit is capable of vapor recompression, such as roots, centrifugal or screw type compressors.

When selecting suitable materials for each unit, specific considerations must be related to the temperature, pressure and chemical and physical properties of the gases and liquids to be treated. However, such considerations will be within the knowledge of those skilled in the art.

Furthermore, one skilled in the art will readily recognize that the selection and control of process parameters will depend on the chemical composition of the gas entering the apparatus and the chemical composition and physical conditions of the gas and liquid in each process step.

The term "reduction of the total energy consumption" is to be understood broadly. According to the invention, a reduction in the total energy consumption means a reduction in the actual energy supplied and/or an increase in the amount of available energy recovered. For the relevant section of a larger facility, the reduced energy consumption can be observed in isolation. However, for such isolated observed sections, reduced energy consumption should not come at the expense of increased energy consumption in upstream and/or downstream sections of a larger facility.

The consumed energy may for example be in the form of electricity, high pressure steam, low pressure steam and/or hot water for warming purposes. Thus, reducing the amount of at least one of these forms of energy in a section of a larger facility is understood to reduce the overall energy consumption, as the case may be.

The following detailed exemplary compositions of the streams are equally applicable to all embodiments of the present invention.

In one embodiment using MEA as absorbent, stream L0 comprises about 70-85 mol% water, about 5-15 mol% MEA and about 2.5-10 mol% carbon dioxide and has a temperature in the range of about 45 ℃ to 50 ℃, e.g., about 47 ℃, and a pressure of about 3 bar.

Stream L1 comprises about 70-85 mol% water, about 5-15 mol% MEA and about 2.5-10 mol% carbon dioxide and has a temperature in the range of about 100 ℃ to 110 ℃, for example about 104 ℃, and a pressure of about 3 bar.

Stream L2 comprises about 80 to 90 mol% water, about 5 to 15 mol% MEA and about 0 to 5 mol% carbon dioxide, and has a temperature of about 105 ℃ to 115 ℃, e.g. about 112 ℃, and a pressure of about 1.4 bar.

Stream L2' comprises about 80 to 90 mol% water, about 5 to 15 mol% MEA and about 0 to 5 mol% carbon dioxide, and has a temperature of about 105 ℃ to 120 ℃, e.g. about 113 ℃, a pressure of about 1.4 bar and a molar fraction of vapour of about 0.1.

Stream L3 comprises about 80 to 90 mol% water, about 5 to 15 mol% MEA and about 0 to 5 mol% carbon dioxide, and has a temperature of about 110 ℃ to 120 ℃, e.g., about 113 ℃, a pressure of about 1.4 bar, and a vapor mole fraction of about 0.

Stream L4 comprises about 99.8 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 0.1 mol% carbon dioxide, and has a temperature of about 70 ℃ to 80 ℃, e.g. about 75 ℃, and a pressure of about 3 bar.

The heated stream L4' comprises about 99.8 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 0.1 mol% carbon dioxide, and has a temperature of about 90 ℃ to 100 ℃, e.g., about 94 ℃, and a pressure of about 1.4 bar.

Stream L4 "comprises about 99.8 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 0.1 mol% carbon dioxide, and has a temperature of about 85 ℃ to 95 ℃, e.g. about 89 ℃, and a pressure of about 0.7 bar.

Stream L4' "comprises about 99.8 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 0.1 mol% carbon dioxide, and has a temperature of about 85 ℃ to 95 ℃, e.g. about 89 ℃, and a pressure of about 0.7 bar.

Stream L4 "" comprises about 99.8 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 0.1 mol% carbon dioxide, and has a temperature of about 70 ℃ to 80 ℃, e.g. about 75 ℃, and a pressure of about 0.4 bar.

Stream L5 comprises about 99.8 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 0.1 mol% carbon dioxide, and has a temperature of about 70 ℃ to 80 ℃, e.g. about 75 ℃, and a pressure of about 0.4 bar.

Stream L6 comprises about 99.8 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 0.1 mol% carbon dioxide, and has a temperature of about 45 ℃ to 55 ℃, e.g. about 50 ℃, and a pressure of about 1.4 bar.

Stream G1 comprises about 55 to 65 mol% water and about 35 to 45 mol% carbon dioxide and has a temperature of about 90 ℃ to 120 ℃, e.g. about 94 ℃, and a pressure of about 1.4 bar.

Stream G2 comprises about 90 to 100 mol% water, about 0 to 2 mol% MEA and about 2 to 10 mol% carbon dioxide, and the temperature is about 105 ℃ to 120 ℃, e.g. about 113 ℃, and the pressure is about 1.4 bar.

Stream G3 comprises about 25 to 35 mol% water and about 65 to 75 mol% carbon dioxide and has a temperature of about 70 ℃ to 80 ℃, for example about 76 ℃, and a pressure of about 1.4 bar.

Stream G4 comprises about 97 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 3 mol% carbon dioxide, and the temperature is about 85 ℃ to 95 ℃, e.g. about 89 ℃, and the pressure is about 0.7 bar.

Stream G4' comprises about 97 to 100 mol% water, about 0 to 0.1 mol% MEA and about 0 to 3 mol% carbon dioxide, and the temperature is about 160 ℃ to 180 ℃, e.g. about 170 ℃, and the pressure is about 1.4 bar.

Stream G5 comprises about 5 to 10 mol% water and about 90 to 95 mol% carbon dioxide and has a temperature of about 45 ℃ to 55 ℃, e.g. about 50 ℃, and a pressure of about 1.4 bar.

Stream a comprises from about 99.8 to 100 mol% water and from about 0 to 0.2 mol% carbon dioxide and has a temperature of from about 70 ℃ to 80 ℃, for example about 75 ℃, and a pressure of about 0.4 bar.

Stream b comprises from about 99.8 to 100 mol% water and from about 0 to 0.2 mol% carbon dioxide and has a temperature of from about 130 ℃ to 145 ℃, for example about 138 ℃, and a pressure of about 0.7 bar.

For all streams, it is suitable that the sum of the constituents does not exceed 100 mol%. The stream may comprise additional components such as nitrogen and/or oxygen.

Examples

The invention is further illustrated in more detail by the following examples. This example should not be construed as limiting the scope of the invention.

Example 1

The following example refers to a calculation example of a section of a larger facility, as substantially shown in fig. 3. The parameters of the streams used in the examples are summarized in table 1 below.

The mild liquid carbon dioxide-rich absorbent stream L0 has absorbed gaseous carbon dioxide, is provided at a mass flow rate of about 16,462kg/h at a temperature of about 47 ℃, and is heated to about 104 ℃, i.e. provides liquid carbon dioxide-rich absorbent L1. Stream L1 then enters a stripping unit (A3) and is mixed with stripper medium G2 supplied at a mass flow of about 1,870kg/h and a temperature of about 97 ℃. The stripping procedure results in a separation into a carbon dioxide-containing gas stream G1 having a temperature of about 94.6 ℃ and a liquid carbon dioxide-depleted absorbent stream L2 having a temperature of about 112 ℃. The carbon dioxide-containing gas stream G1 was cooled by direct contact heat exchange with a stream of liquid heat transfer medium L4 supplied at a flow rate of about 15,000kg/h and an initial temperature of about 76 ℃ to provide a cooled carbon dioxide-containing stream G3 at a temperature of about 76 ℃ and a heated stream L4' at a temperature of about 94 ℃. The heated stream L4' is subsequently depressurized using a depressurization valve (a8) into a depressurized stream L4 ". The depressurized stream L4 "was separated in a flash separation unit (a9) providing recovered stripping medium G4 at a temperature of about 90 ℃ and a pressure of about 0.7 bar, which was compressed in a second pressure increasing unit (a12) forming a temperature of about 170 ℃ and a pressure of about 1.4 bar. The flash separation also produces a liquid stream L4 "' having a temperature and pressure as recovered stripping medium G4. The liquid stream L4 "'is then further depressurized in a second depressurization unit (a13) to a pressure and temperature of 0.4 bar and 76 ℃, the liquid stream L4"' is then separated into a second gas a which is recompressed in a first compressor (a11) to a compressed second gas b to a pressure of 0.7 bar. The second gas is returned to the flash separation unit (a) and exits as part of the recovered stripping medium G4.

The liquid carbon dioxide-depleted absorbent stream L2 leaving the stripper (a2) is heated in a third heat exchanger (a6) using external heat, which provides a gas/liquid heated carbon dioxide-depleted absorbent stream L2 'at a temperature of about 113 ℃ for mixing with the compressed recovered stripping medium G4' in the second separator (a7) and a mass flow rate of about 16,760kg/h (L2 ') and 538kg/h (G4'), respectively. The mixture of gas/liquid heated carbon dioxide-depleted absorbent stream L2 'and compressed recovered stripping medium G4' is then separated, thus providing a recovered liquid absorbent stream L3 with a mass flow rate of about 15,428kg/h, and a stripper medium G2 with a mass flow rate of about 1870kg/h, a temperature of about 113 ℃ and a pressure of 1.4 bar.

As can be seen from this example, the recycle of water from the carbon dioxide-containing gas stream (G1) can be provided to the stripping unit (a2) again and constitutes about 1/3 of the stripping medium without compromising the yield and purity of the carbon dioxide formed.

Thus, one third of the energy additionally supplied in the form of a steam circuit can be replaced by working with a stream of already existing steam.

Based on the above described embodiments, the power absorption of the first pressure boosting unit (a11) is 13kW and the electric power absorption of the second pressure boosting unit (a12) is 23kW, i.e. 36kW for operating these units. This electric power corresponds to 108kW steam power.

The steam (G4') generated by the process in the examples amounted to 334kW (2230kJ/kg 523kg/3600), so the total energy reduction of the illustrated example was 226kW steam. This corresponds approximately to an energy reduction of more than 20%.

Claims

1. A process for recovering acid gas from a liquid acid gas-rich absorbent stream (L1), comprising the steps of:

a. providing the liquid acid gas-rich absorbent stream (L1) having absorbed acid gas therein,

b. separating acid gas from the liquid acid gas-rich absorbent stream (L1) in a stripping column (A2) using a gaseous stripping medium (G2) to provide an acid gas-containing gas stream (G1) and a liquid acid gas-depleted absorbent stream (L2),

c. transferring heat from the acid gas-containing gas stream (G1) to a stream of heat transfer fluid (L4) to provide a cooled acid gas-containing stream (G3) and a heated stream (L4'), wherein the heat transfer is provided as follows: directly contacting the heat transfer fluid (L4) with the acid gas-containing gas stream (G1) to obtain a heated stream (L4') and a cooled acid gas-containing gas stream (G3),

d. separating the heated stream (L4') into recovered stripping medium (G4) and liquid heat carrier (L5), wherein the liquid heat carrier (L5) is used as at least a portion of the heat transfer fluid (L4), and

e. the recovered stripping medium (G4) is provided directly or indirectly to the stripper column (a 2).

2. The method of claim 1, wherein the acid gas is carbon dioxide.

3. Process according to claim 1 or 2, wherein the pressure of the recovered stripping medium (G4) is higher than or equal to the operating pressure in the stripping separation in step b.

4. The process according to claim 1 or 2, wherein the absorbent is aqueous.

5. Process according to claim 1 or 2, wherein the recovered stripping medium (G4), optionally compressed, is fed directly to the stripper column (A2) at a position below the feed position of the gaseous stripping medium (G2).

6. The process according to claim 1 or 2, wherein the recovered stripping medium (G4) is compressed to provide a compressed recovered stripping medium (G4') before being provided to the stripping column (a 2).

7. The method according to claim 1 or 2, wherein the separation of step d.is provided by:

separating the reduced pressure stream (L4 ") by flashing in a first flash column (a9) to provide recovered stripping medium (G4) and liquid heat carrier (L5).

8. The method according to claim 7, wherein the separation of step d.ii. provides a liquid stream (L4 "') which is subjected to the following steps, before providing the liquid heat carrier (L5):

further depressurizing the liquid stream (L4 ') to provide a second depressurized stream (L4 ') having a pressure lower than the pressure of the liquid stream (L4 '),

recompressing the second gas (a) to provide recompressed second gas (b); and

d.vi. feeding the recompressed second gas (b) to the first flash column (a9) where it leaves the first flash column (a9) as part of the recovered stripping medium (G4).

9. The method of claim 8, wherein all steps d.i. to d.vi.

10. The method of claim 9 wherein all steps d.i. to d.vi.2, 3 or 4 are repeated.

11. The method according to claim 9 or 10, wherein the repetition is sequential and/or parallel.

12. Process according to claim 1, wherein the heat transfer fluid (L4) is depressurized to a pressure lower than the pressure of the liquid heat carrier (L5) before heat transfer.

13. The method according to claim 1 or 2, wherein the separation of step d.is provided by:

separating the heated stream (L4 ') by flashing in a first flash column (a9) to provide the recovered stripping medium (G4) and a liquid stream (L4 "');

separating the second reduced-pressure stream (L4 "") in a third separation unit (a10) to provide a second gas (a) and the liquid heat carrier (L5);

recompressing the second gas (a) to provide recompressed second gas (b); and

v. feeding the recompressed second gas (b) to the first flash column (a9) where it leaves the first flash column (a9) as part of the recovered stripping medium (G4).

14. The process according to claim 1 or 2, wherein the acid gas containing stream (G1) is compressed into a compressed acid gas containing stream (G1') before the heat transfer step.

15. The process of any one of claims 8-10 or 13, wherein any one or more of the heated stream (L4 '), the depressurized stream (L4 "), the liquid stream (L4"'), the second depressurized stream (L4 "") is heated by a heat source.

16. The process according to claim 15, wherein the depressurized stream (L4 ") and/or the second depressurized stream (L4" ") is heated by a heat source.

17. The method according to claim 6, wherein the method further comprises the steps of:

f. heating the liquid carbon dioxide-lean absorbent stream (L2) to provide a gas/liquid heated carbon dioxide-lean absorbent (L2');

g. the gas/liquid heated carbon dioxide lean absorbent (L2 ') is separated in a second separator (a7) to provide an evaporated stripping medium (G2').

18. The process according to claim 17, wherein the stripping medium (G2) is an evaporated stripping medium (G2').

19. The process according to claim 18, wherein the stripping medium (G2) comprises the compressed recovered stripping medium (G4').

20. The process according to claim 19, wherein the compressed recovered stripping medium (G4') and the vaporized stripping medium (G2') are mixed before being fed to the stripping column (a 2).

21. The method according to claim 1 or 2, further comprising the step of:

-optionally cooling the cooled acid gas containing stream (G3);

-separating the cooled acid gas-containing stream (G3) into an acid gas product stream (G5) and a second liquid stream (L6), optionally further cooling the cooled acid gas-containing stream (G3) beforehand, and

-optionally recycling the second liquid stream (L6) to the stripper column (a 2).

22. The process according to claim 1 or 2, wherein the liquid acid gas-depleted absorbent stream (L2) is heated in a third heat exchanger (a6) and separated to provide the gaseous stripping medium (G2) and a recovered liquid absorbent stream (L3), and wherein the heat transfer fluid (L4) of step c is a recovered liquid absorbent stream (L3).

23. The process according to claim 22, wherein the recovered liquid absorbent stream (L3) is depressurized to provide a depressurized recovered liquid absorbent stream (L3 '), the depressurized recovered liquid absorbent stream (L3') is heated in a heat recovery unit (A3) to provide a heated recovered liquid absorbent stream (L3 ").

24. The process according to claim 23, wherein the heated recovered liquid absorbent stream (L3 ") is separated in a fourth separator (a19) to provide absorbent and the recovered stripping medium (G4).

25. The process according to claim 24, wherein the recovered stripping medium (G4) is pressurized in a second pressurization unit (a12) to provide a compressed recovered stripping medium (G4'), the compressed recovered stripping medium (G4') being provided to the stripping column (a 2).